Process for refining high sulfur crude oils



June 17, 1952 w, SLATER PROCESS FOR REFINING HIGH SULFUR CRUDE OILS Filed Aug. 29, 1950 2 SHEETSSHEET 1 09H QUQKOHM OH *0 INVENTOR.

A. SLfl'IER ILLIAM ATTORNEZY June 1952 w. A. $LATER 2,600,931

PROCESS FOR REFINING HIGH SULFUR CRUDE OILS Filed Aug. 29, 1950 2 SHEETSSHEET 2 ,7 ilf Q m i 5? O 4- 0 INVENTOR. ILLIAM fi- Lfi E R w ATTQRNEY is because the gasoline is the product of primary interest. However, in many cases it may also be desirable to remove materials boiling higher than gasoline, such as furnace oil, with the remainder or a suitable cut thereof being sent to the cracking step.

In the following description certain preferred modifications of the invention have been described. It is understood that these are by way of illustration only and are not to be considered as limiting.

Referring briefly to the attached drawings, Figure 1 presents a diagrammatic arrangement of a suitable apparatus for carrying out my two-stage process. Figure 2 is a graphic plot of depth of conversion against coke laydown for a cracking reaction carried out on charge stocks which have been hydrodesulfurized catalytically and by the absorption process.

In the operation of my process I charge a residual-oil-containing charge such as crude to an absorption hydrodesulfurization unit. Absorption hydrodesulfurization may be distinguished from catalytic hydrodesulfurization in that the former employs a catalyst which chemically reacts with or absorbs sulfur contained in the charge, whereas the latter employs a catalyst which is relatively immune to the action of sulfur and which causes conversion of sulfur in the charge to hydrogen sulfide. sorption process is terminated before substantial amounts of hydrogen sulfide occur in the efliuent, while the catalytic process inherently necessitates the presence of hydrogen sulfide in the effluent throughout the entire operation.

The charge is treated with a suitable sulfur absorbing contact, e. g., nickel, nickel oxide, or partially reduced nickel oxide, at elevated temperature and pressure in the presence of hydrogen. During the treating process the sulfur in the charge stock reacts with the sulfur absorbent such as nickel or nickel oxide, to form solid nickel sulfide. Hydrogen replaces the sulfur in the hydrocarbon radical thus produced. The sulfur atoms which are replaced by hydrogen may exist in various places in the hydrocarbon molecule. Where the sulfur atom acts as a link between two hydrocarbon radicals, replacement of the sulfur atom with hydrogen may prevent recombination of the molecular fragments, thus produced. The cracking which takes place may be thermal, catalytic or some of both. However it should be observed that the cracking which occurs in the process does so with a minimum formation of condensed rings and high molecular weight materials (coke or heavy bottoms) which usually accompany ordinary cracking. Another reaction which takes place to a certain extent during the contact period is the saturation of unsaturated linkages in the original or cracked hydrocarbon molecules.

As more charge stock is passed through the catalyst bed, the catalyst particles absorb more and more sulfur. When most of the nickel or nickel oxide has been converted to nickel sulfide,

substantial amounts of hydrogen sulfide will begin to appear in the effluent. The processing period is terminated before this point'and the charge may be switched to another reactor which is ready to go on-stream. The catalyst in the first reactor is regenerated by passing air Also the ab- (diluted with steam or inert gas as desired) therethrough. When most of the sulfur and carbon have been burned ofi, the contacts are ready for reuse in the processing cycle described above.

Following the absorption hydrodesulfurization stage, the hydrodesulfurized crude is topped to remove at least gasoline and lighter constituents, and in some instances the hydrodesulfurized crude may be topped to remove constituents somewhat heavier than gasoline, say the lighter materials down through furnace oil. ..In either event the remaining material, heavier than gasoline, or a suitable cut of the remainder, is then subjected to a catalytic cracking treatment of known type. The cracking charge is cracked to a depth of conversion above about 30 weight per cent.

The depth of conversion is determined by the conditions under which the process is operated. The conditions necessary to bring about a desired degree of conversion of a given charge stock are obtained by experience. The per cent conversion at given conditions and for a given charge stock may be predetermined by experiment. Since per cent conversion refers to the total gasoline, gas, and coke produced in the process, the determination of conversion depth involves measurement of the amounts of gasoline, gas, and coke formed. In laboratory work, depth of conversion is usually given as weight per cent, while in commercial operations it is usually given as volume per cent. Since the values given in either form are approximately equal, the per cent conversion as used herein and in the claims is meant to include either weight or volume per cent.

Referring now to Figure 1 in detail, in order that my process may be more clearly understood, fresh charge, for example crude oil, is introduced into the system through line '24. The charge passes from line 24 into heater 26, where it is heated to reaction temperature, or possibly somewhat lower, to allow for the temperature increase across the contact bed. The preheated charge then passes into transfer line 28.

Fresh hydrogen enters the system through lines 2 and 3 and is heated in heater 4 to approxi- 'mately the same temperature as the charge stock.

The preheated hydrogen passes through line 6 to join the preheated charge from line 28. From line 6 the mixture of heated hydrogen and charge stock passes through line 8, through valve 10, through line l2 into reactor l4.

Within the reactor the sulfur in the charge is chemically combined with the metalliferous contact agent. The cracking and hydrogenating reactions described previously also take place.

Converted product passes out of reactor [4 into line 46, through valve 48, into line 50 and into cooler 52, where the products are partially cooled.

When reactor [4 is off-stream reactor 20 is on a processing cycle. In this instance preheated hydrogen and charge stock pass from line 6 through valve I3, into line I6, through line I8 and into reactor 20, where the reactions described above occur. Converted products pass out of reactor 20 through line 54, through valve 56, into line 50 and into cooler 52.

From. cooler 52 the partially cooled hydrodesulfurized charge passes through line 10 into separator 12, where an initial separation takes place. Hydrogen dissolved in the product and possibly some gaseous hydrocarbons are flashed off through line 13 into absorber l6. Gaseous line II 6 as stabilizer bottoms.

hydrogen and whatever light hydrocarbons are mixed therewith pass out of absorber 16 through line I8 into line 80, from whence they are recycled into the system via line 3 and heater 4. A portion of the gases from line I8 may be bled out through valve I62, line I00, into line 94, through valve 96, into line 98 and vented, in order to prevent the building up of inert materials, e. g., methane, in the recycle hydrogen.

The liquid product in separator I2 passes through line 82 into fractionator 84. Light hydrocarbon gases pass out of the fractionator through line 94 as overhead. From line 94 these gases pass through valve 96 and into line 96 from which they may be vented. The gasoline fraction is removed from fractionator 84 through line I04 into stabilizer I06. C4 hydrocarbons pass out of the stabilizer as overhead through line I66 and thence to storage. Hydrodesulfurized straight run gasoline is removed through The straight run gasoline referred to in the accompanying description is the gasoline obtained from the hydrodesulfurized product Without further cracking. While it is not purely straight run gasoline in the normal sense in view of the pretreating step, it is essentially straight run in character,

being predominantly a saturated type of gasoline.

Bottoms from fractionator 84 pass through line 86 to charge drum H2 and thence to the second stage of my process. A portion of these bottoms are diverted from line 86 and recycled through line 88, through pump 96, into line 92 and into absorber 16.

impurity, however, render the gasoline unsatisfactory for many purposes and require removal.

It is advantageous to conduct a hydrogen purge of the deactivated catalyst in the offstream reactor prior to regeneration of the catalyst. This may be accomplished by passing recycle and/or fresh hydrogen through lines 6 and I to line 9. through valve II, through line I5. through line 36, valve 38, line 46. line I2 and into reactor I4. The purge hydrogen at -reaction pressure and at about reaction temperature acts to sweep out the converted product remaining in the reactor and also effects a partial conversion of the carbonaceous deposit on the catalyst, thus increasing the total yield and allowing easier regeneration. The purge hydrogen leaves reactor I4 through line 46, through valve 48 into line 50 from which it follows the course of the hydrodesulfurized charge stock. While a hydrogen purge is' considered desirable, it is by no means necessary to the invention, since the desirable results of my process are brought about with or without a hydrogen purge.

At the end of the hydrogen purge valve 43 is closed and the reactor is depressured through line 46, valve 66 and line 62.

At the completion of the hydrogen purge a steam purge is considered advantageous in order to remove whatever traces of hydrogen or re- 30, heater 32, and into line 34.

actants may remain in the reactor. This may be accomplished by passing steam through line From line 34 the steam passes through line 36, valve 36 and line '46, through line I2 and into reactor I4. Purge steam leaves the reactor by way of line 46, line 58, valve 60 and into line 62 to vent.

At the completion of the steam purge, air or other oxygen-containing gas (which may be diluted with steam or other inert gas such as flue gas) is passed through line 36, through heater 32 and into line 34. From line 34 the regenerating gas passes through line 36, valve 38, line 46, line I2 and into reactor I4. The regenerating gas operates to burn 01f the contaminants on the metalliferous contact in the reactor. Flue gas leaves reactor I4 by way of line 46, line 58, valve 60, and line 62. From line 62 the flue gas may be vented. Alternatively, the flue gas may be passed through a scrubber (not shown) to remove sulfur dioxide and recycled as a diluent into the regenerating gas entering the system through line 36.

When reactor 20 is cit-stream, purge hydrogen passes from line 6, through line I, through valve II, through line I9, valve 42, lines 44 and I3, and into reactor 20. Purge hydrogen and purged product leave reactor 26 through line 54 and valve 56 and pass into line 56, from which these materials follow the course described for the hydrodesulfurized charge stock. Reactor 26 is depressured through line 54, line 55, valve 66 and line 62. Purge steam passes from line at through valve 42, through line 44, through line I8 and into reactor 29. Steam and purged products are vented through line 54, line 6 5, through valve 68, and into line 62. Regenerating gases flow from line 34 into reactor 26 following the same course as that described for the purging steam.

The topped hydrodesulfurized crude passes from charge drum II 2 through line H4 into heater H6, where it is heated to cracking temperature. From the heater the preheated cracking stock passes through line H6 into lines I23 and I22 and into the catalyst bed in cracking reactor I26.

Although a Thermofor catalytic cracking unit has been illustrated in the drawing it will be understood, of course, that any other conventional type of catalytic cracking apparatus, employing a fluidized, fixed or moving bed, may be employed.

Active catalyst is continuously supplied to re actor I26 from the top of catalyst elevator I2 3 by means of catalyst transfer line I28. Deactivated cracking catalyst leaves reactor 26 continuously by way of transfer line I! and conveyed from the bottom to the top of catalyst elevator I32. From the top of elevator I32, the deactivated catalyst passes to regenerator by way of line I34. An oxygen-containing gas such as air is continuously introduced into regenerator I36 by means of lines I 46 and I42 in order to burn oiT the coke which has been deposited on the catalyst. Flue gas passes out of regenerator I36 through lines I 4| and I43 The exothermic heat of the regenerating reaction is removed by circulating cooling fluid such as water through lines I44 and I46. The water takes up the heat of regeneration and leaves the regenerator as steam through lines I45 and MT. Hot regenerated catalyst leaves regenerator I35 at its lower end and passes through line I36 to the lower end of catalyst elevator I24. The hot regenerated catalyst is then again passed from the top of catalyst elevator I24 through line I28 and into reactor I26. Hydrating steam may be supplied to the cracking reactor I25 through line I48.

The products of the cracking reaction which takes place in reactor I26 pass through line I49 into fractionator I59. Hydrocarbon gases leave the fraetionator as overhead through line I52 and bottoms are removed through line I55. Intermediate fractions, including gasoline, are removed through lines I56, I58 and I60.

It will be understood, of course, that the attached drawing is purely schematic, consequently, all valves, pumps, recycle drums and like details have not been shown. However, the equipment illustrated in the drawing is entirely conventional and omitted details may be readily supplied by one skilled in the art.

As mentioned previously, the utilizing of absorption hydrodesulfurization prior to cracking has among its unique achievements those most important advantages of lower coke laydown on the cracking catalyst and improved gasoline yield. In order to show these and other advantageous results more clearly I have presented below data comparing the results of absorption hydrodesulfurization following by cracking with results obtained by catalytic hydrodesulfurization followed by cracking. Table I presents the laboratory inspections on the charge. The lefthand column gives information on the charge which was hydrodesulfurized by the absorption method in run A. The right-hand column describes the charge used in run B, a catalytic hydrodesulfurization.

Table II illustrates processing conditions, catalysts, and adjusted yields for the two types of hydrodesulfurization. Run A deals with absorption hydrodesulfurization and run B deals with catalytic hydrodesuliurization.

TABLE II Run A 13 Charge Whole, De- West Texas salted. Crude. Catalyst 20% Ni (as 12% Ni-W NiO) on a (1:1) as support. NiO W03 on a support Processing Conditions:

Reactor Inlet Temperature, F.. 840 841. Reactor Outlet Temperature, {F 850 843 Average Reactor Pressure, p.s.1.g. 510 520 Process Period Length, hr 3. 4. 0 Gas Purge Period Length, Hn- 2.0 1. 42 Charge Rate, BbL/Hr 2.025 2. 1534 Liquid Hourly Space Velocity,

Vol. Oi1/Hr./Vol. Cat 1.01 1. 067 Gas to Reactors, Rate, SCF/Bbl.

Charge 9, 724 7, 210

TABLE IICont1nued Run A B Average H2 Content, Mol Percent. 85. 6 70. 9

Hydrogen Consumption, 50

Bbl. Charge 340 426 Adjusted Yields Volume Percent Charge:

Total C4 2. 41 7. 25 Total 5 2.07 6. 76 Depentanizcd Total Liquid Pro- Total 97. 37 96. 62 Weight Percent Charge:

Dry Gas, Ca and lighter 1. 90 4.31 otal C l. 64 4. 94 l. 58 4. 99

Table III presents the laboratory inspections on the total products from the two types of hydrodesulfurization.

TABLE III Run A 13 Charge Whole, De- West Texas salted. Crude. Catalyst i- 20% Ni (as 12% N i-W N10) ona (1:1) (as support NiO WOa) on a support. Total Product: Gravity, API 39. l 43. 4 Color, NPA 7. 25 4. 75 Sulfur, B, Percent 0. 45 0.18 Carbon Residue on Btms. Above 590.F. V.T., Percent. 0.88 0.35 Distillation, Gas Oil ASTM Dl58-41:

Over Point, "F 140 Percent at 392 F. 3. 42 500 48 G2 590 62 76 Regarding the significance of Tables I, II and III, both runs A and B were carried out on a whole desalted West Texas crude of substantially identical properties and under substantially the same conditions. The absorption hydrodesulfurization (run A) employed a nickel oxide catalyst, while the catalytic hydrodesulfurization (mm B) utilized a nickel tungstate, sulfur-immune catalyst. The catalyst in each run was supported on a commercial carrier. It should be noted that the catalyst carbon in the absorption process was 0.95 per cent by weight of charge, whereas in the catalytic process it was 1.67 per cent by weight of charge.

Table IV presents conversion data obtained by cracking the hydrodesulfurized crude obtained from test runs A and B illustrated in Tables I, II

and III. r

The hydrodesulfurized crudes in each case were topped to remove gasoline and lighter prior to the cracking treatment. Eight runs on the topped crudes were made, three on the topped catalytically hydrodesulfurized crude and five on the absorption hydrodesulfurized crude. The runs were made at comparative space velocities for each charge, with other conditions being substantially identical.

TABLE IV Conversion indew tests on. topped hydrodesuljurized West Texas crude Weight West Texas Crude 1 C Temo, Percent Percent Percent Percent Hydrodesulfurized over F. LHSV P i3 gags; Coke Cgligler N i-W-on-support (from Test 1 851 2 0. 5 l9. 7 8. 2 3. 2 27. 9 Run B) 2 854 1 0. 5 22. 3 11. 8 4. 2 34.1 3 854 0. 5 0. 5 22. 9 18. 5 6. U 41. 4 4 849 2 0. 5 23. 8 9. 2 3. 9 33. O N1-O-ou-support (from Test 5 852 1 0. 5 25. l 14. 3 4. 6 39. 4 R1111 A) 6 850 2 n. 5 23. 4 10. 1. [1 3 3. 4 7 856 0. 0. 5 25. 3 19. 5 5. 8 .41. 8 8 S57 0. 5 0. 5 25. l 18. 7 5. 6 43. 8

1 Topped after hydrodesulfurization to remove 400 F. EP gasoline before cracking.

Figure 2 is a semilogarithmic graph plotting reduced hydrodesuliurized West Texas crude obthe results shown in Table IV and, more specifitained by the absorption hydrodesulfurization cally, plotting the weight per cent conversion for method. The information under runs 3 and 4 each type of charge against the weight per cent represents the results obtained from fluid cataof coke. It will be seen that above about lytic cracking of reduced. catalytically hydrodeweight per cent conversion (30 weight per cent sulfurized West Texas crude. The charge in runs gasoline, gas and coke) the carbon laydown is 1 and 2 was obtained by topping thep'roduct obless for the cracking stock which has been hydro tained from run A (Tables II and III) to remove desulfurized by the absorption method (curve A) constituents boiling below about 625 F. The than for the catalytically hydrodesulfurized charge stock of runs 3 and 4 was obtained by recracking stock (curve C). This effect is totally moving the front ends of the product from run unpredictable as will be seen from reference to B (Tables II and 111) $0 e e p th Table HI. This table shows that the carbon Cracking char mployed in runs 1-4 therefore residue on bottoms was 0.88 per cent for the abg remaining bottoms boiling above about $8??? .iei fifi iii fiitttififit fiyiiiiti Table v describes the e ee e p furized product, or more than twice as great. r a d charges The dlstmaffimn data m thls In order to bring out the full import of the difw table mdmates that t absorptmn. hydroqesul' ference in coke laydown produced by the two Tig g gi g g f g Y i orthlgher types of cracking stocks it should be pointed out ggg 52 2 3 .3. 25 g g iid gfig that at per cent conversion, for example, the treated charge du ed a, higher percentage of difierence in coke deposited is about 0.7 per cent. 0 gasoline and laid down much less coke This is When char in 0'b of cracking Stock additionally unexpected in view of the higher per day (app b y 30.0 po n p bb carbon residue on the cracking charge obtained 6,000,000 pounds-charged) this difference is equal by absorption hydrodesulfurization.

TABLE V Run Number l 2 3 4 Reduced Hydrode- ReducedHydrodesulsuliurized West iurized West Texas Aniline Point, F; Carbon Residue, Weight percent.. Sulfur, Weight percent Distillation, 'F. (Vacuum Corrected to 760 mm): lnitilrg Boiling Point 1 Same as for run 1. 2 Same as for run 3.

to about 0.007 6,000,000 pounds or 21 tons less Table VI presents data on catalyst, operating coke to be burned ofi daily. conditions and estimated commercial operating Several fluid cracking runs were also made on conditions. From this table it will be seen that topped crudes' which had been hydrodesul-iuri-zed run 1 is comparable with run 3, since both were by the catalytic and absorption method-s. Tables made at a conversion depth of about 58 weight V, VI and VII present the resultsin these runs. per cent. Similarly, run 2 is comparable with The informationunder runsl and 2 represents run 4, since both runs. were made at a conversion the data obtained from fluid catalytic cracking of depth of about 68 weight per cent.

' TABLE VI Run Number 1 2 3 4 Charge Stock Reduced Hydrodesuli'ur- Reduced Hydrodesulfurized West Texas Crude ized. West Texas Crude (over supported NiO) (Produced over s up ported Nickel Tungstate Catalyst) Cat Natural, SR Filtrol Natural, SR Filtrol 23. 4 23. 4 23. 6 23. 6 Carbon Factor 1.15 1.15 1.12 1.12 Operating Conditions:

Average Reactor Termperature, "F 926 925 925 925 Average Reactor Pressure, p.s.l.g- 12.8 12.9 13. 2 13.0 Outlet Cycle O11 Partial Pressure, p. s. 1. a... 5. 2 5. 9 5. 6 6.5 Space Velocity, wt./hr./wt. (Basis Total Feed) 3. 3 3. 7 1. 7 1. 9 Catalyst to Oil Ratio, wtJwt. (Bas s Total Feed) 10. 4 10.0 9. 5 10.1 Catalyst to Oil Ratio, wt/wt. (Basis Fresh Feed) 14.4 18.1 13.1 18. o Recycle, volume percent iresh feed 36. 6 78. 2 25. 8 78. 4 Furnace Oil 0. 0 0. 0 0. 0 0.0 34. 1 75. 5 33. 3 75. 9 yOil 2.5 2.7 2.5 2.5 Dispersion Steam, lb./1000 lb. eatalyst.. 1. 7 1. 8 1. 9 1. 8 Stripping Steam, lb./1000 lb. catalyst 5. 8 6. 0 6. 3 6. 0 Stripping Nitrogen as Steam, lb./1000 lb. catalyst. 0. 0 0. 0 0. 0 0. 0 Rehydration Steam, lb./1000 lb. catalyst 0. 0 0. 0 0. 0 0. 0 Total Reactor Inerts as Steam, weight percent of fresh food 8. 5 11. 0 8.5 11.4 Carbon on Spent Catalyst, percent on catalyst 0.635 0.662 1.041 0.1171 Carbon on Regenerated Catalyst, percent on catalyst 0. 173 0.192 0.209 0.196 Conversion, volume percent of fresh iced 60.0 69. 3 59. 7 71. 6 Conversion, weight percent 58. 8 07. 8 57. 8 09. 0 Estimated Commercial Operating Conditions:

Outlet Cycle Oil Partial Pressure, p. s. i. a 5. 4 G. 3 5. 9 6. 9 Catalyst to Oil Ratio, wt./wt. (Basis Total Feed)... 10.0 10.0 10.0 10.0 Carbon on Spent Catalyst, percent on catalyst... 1.138 1.131 1. 464 1. 458 Carbon on Regenerated Catalyst, percent on cataly 0.600 0.600 0.600 0. 600 Carbon Burnofl, lb./hr./l000 bbl. iresh feed/day..- 982 1, 274 1, 593 2, 111 Carbon Burnofl, lb./hr./1000 bbl. gas oil converted/day 1,636 1,838 2, 71 2, 949

Table VII presents yield data. from runs 1 through 4 corrected to estimated commercial opcoke laydown is even more striking in these runs than in Table IV.

TABLE VII Run Number 1 2 3 4 Charge Reduced Hydrodesuliur- Reduced Hydrodesullur- Y ized West Texas Crude, izcd West Texas Crude over supported NiO (Produced over supported Nickel-Tungstate Catalyst) Catalyst Natural, SR Filtrol Natural, SR Filtrol Yields, volume per cent of fresh feed (Corrected to 100% Weight Balance): Debutanized Motor Gasoline 44.1 48. 9 37. 5 42. 6 Isopentane 2. 4 2. 7 3. 4 3. 9 N-Pentane 0. 7 0.6 0.8 0.7 Pentenes 5- 8 6. 5 4. 1 4. 8 Hexanes and Heavier. 35. 2 39.1 29. 2 33.2 Butane-Butene 9. 8 11. 8 l0. 4 12. 5 Is0butane. 2. 8 3. 5 3. 7 4. 8 N-Butane- 1.0 1. 1 1.4 1. 4 Butene 6.0 7. 2 5. 3 6. 3 Propane-Propylene. 7. 0 8. 3 0. 9 8. 8 ropane 2. 2 2.6 2.5 3. 2 Propylene 4. 8 5. 7 4. 4 5. 6 Furnace 011.... 14. 7 18. 2 12.4 12.8 Heavy Cycle 11. 20. 2 7. 5 23.0 10.5 Slurry 011.... 5. 1 5.0 4. 9 5.1 Total 100. 9 99. 7 95. 1 92. 3 10 RVP Motor Gasoline 48. 4 54.0 40. 5 45. 8 Excess Butane-Butene. 5. 5 6. 7 7. 4 9. 3 Gas (04 and Lighter), weight per cent 13.6 16.1 13. 8 17. 1 Gas (01 and Lighter), weight per cent 3. 2 3. 8 3. 3 4. 1 Coke (0.9 Carbon, 0.]. Hydrogen), weight per cent.. 8. 3 10. 7 13.3 17. 5 HZS, weight per cent 0.3 0.4 0.1 0.2 Depropanized Gasoline, bbl./1000 bbl. 011 converted. 898 876 802 770 Yields corrected to estimated commercial operating conditions.

crating conditions. It should be noted that run 1 produced substantially more gasoline than run 3 and that run 1 produced substantially less coke and gas than run 3. The same is true of Table VIII combines portions of the data shown in Tables II and III with data from fluid cracking runs 1 and 3 of Tables V and VIII. The data presentedhere illustrate the improvement in gasorun 2 as compared with run 4. The difference in 75 line yield to be obtained by my process.

TABLE v11:

Hydrodesuliurization Run No.

Charge Hydrodesulfurization Catalyst Total Liquid Product, Volume Per Cent of West Texas Crude Gasoline, Volume Per Cent of Total Liquid Product. Gasoline, Volume Per Cent of West Texas Crude Fluid Cracking Run No Charge, Volume Per Cent of Total L1quid Hydrodesulfurized Produ Conversion, weight Per Cent of Fresh Feed" Gasoline (1O RVP), Volume Per Cent of Charge Gasoline (10 RVP), Volume Per Cent of Total L desulfurized Product Gasoline (l RVP), Volume Per Cent of West Texas Crude Total Gasoline, Volume Per Cent of West Texas Crude Difference in Gasoline Yield, Volume Per Cent of West Texas Crude BblsJday, (20,000 bbl. unit) (A mmus B) West Texas West Texas Crude Crude Supported Supported NiO N ickel-Tungstate 97. 37 96. 62 35 42 34. 0 40.

1 Per Cent 392 F. by ASTM D-158-41 Distillation.

In order to illustrate that the advantages of my process are not limited merely to fluidized cracking, data obtained from the fixed-bed cracking of reduced hydrodesulfurized West Texas crude are presented in Table IX. The cracking stock employed in run 5 was identical with that used in runs 1 and 2 (Table V). The cracking stock of run 6 was identical with that employed in runs 3 and 4 (Table V) It will be noted that runs 5 and 6 of Table IX were carried out to approximately identical depths of conversion. The coke on catalyst was substantially lower in run 5 than in run 6, again despite the fact that carbon residue for the cracking charge of run 5 was higher than that for run 6.

TABLE IX Reduced hydrodesulfurized West Texas crude Run Number 5 6 2A9 Agtvlty Stealfi l ge ranu ar cm W Filtrol Operating Conditions:

' Catalyst Temperatures, F.:

Average Crackmg 924 Average Maximum Regeneration. 997 Charge Period, minutes Space Velocity, wt./ hr./wt Catalyst-to-Oil Ratio, wt./wt Dispersion Steam, lb./1000 lb. catalyst. Rehydration Steam, lb./1000 lb. catalyst Carbon Burnofi, Weight per cent on catalyst/cycle Carbon Burnofi, lb./hr./l000 bbl. o1l/day Conversion, volume per cent on charge. Conversion weight per cent on charge Yields, volume per cent on charge (Corrected to 100% Weight Balance):

Debutanized Motor Gasoline lsopentane N-Pentane- Pentenes Hexanes and Heaviei Butane-Butene 27-683 :oqocoom N-Butenes Propane-Propylene. 8. 8

Catalytic Gas Total Gas (04 and Lighter), weight per cen charge Gas (02 and Lighter), weight per cent on charge Coke (0.9 Carbon, 0.1 Hydrogen), welght per cent on charge H28, weight per cent on charge The data set forth in Tables I-Di clearly illustrate the unique benefits of absorption hydrodesulfurization in the two-stage process. Coke laydown on the cracking catalyst is reduced substantially. Furthermore, a greater yield of gasoline is realized. These advantages are accomplished with the concurrent reduction in sulfur contained in the virgin charge.

Since the first stage of my process (absorption hydrodesulfurization) is concerned with treatment of charge stocks containing residual constituents which require extensive conversion, I prefer to employ rather high temperatures desirably between about 750 and 950 F., and particularly in the range of about 800-870 F., although lower or higher temperatures may be employed. Pressures in the hydrodesulfurization stage may vary between about and 1500 p. s. i. g. Pressures of about 500 to 1000 p. s. i. g. are preferred. Cost of equipment is kept low, and hydrogen consumption is quite low at these relatively low pressures. The hydrogenzoil ratio during the hydrodesulfurization reaction may vary quite widely between about 1000 and about 20,000 cubic feet of hydrogen per barrel of oil and preferably between about 5000 and 12,000 ft. I-Iz/bbl. oil. Space velocities may vary between about 0.2 and about 6 and preferably between about 0.5 and 2.0 liquid volumes of oil per volume of catalyst per hour. Throughput advantageously varies between about 1 and about 10, with a throughput of about 2-8 being very satisfactory in most cases. The space velocity and throughput vary according to the type of charge stock, i. e., its content of heavy constituents, and the degree of desulfurization and conversion desired. I have found that any hydrogen purity above about 50 per cent is satisfactory.

The hydrodesulfurization catalyst employed in my process may be any iron group metal (iron, cobalt and nickel), iron group metal oxides, or combinations of one or more of these metals and/ or their oxides, desirably deposited on a porous support or carrier. Nickeliferous catalysts are preferred.

When employed in conjunction with a carrier the contact agents may be made by impregnation of the porous support or carrier with a solution of a soluble salt, such as a nitrate, followed by calcining to form the oxide, and followed by reduction if a metal or a mixture of metal oxides is to be used. Relatively large amounts of metal or metal oxides should be deposited on the carrier and preferably from about 2 to about 30 per cent. Examples of suitable porous carriers are activated alumina, silica-alumina cracking catalysts of conventional type prepared synthetically or by acid treatment of clays, silica gel, Magnesol, Porocel, Alfrax, kieselguhr and others.

My invention is applicable to hydrocarbon oil charge stocks containing residual constituents. By residual constituents I mean difliculty vaporizable hydrocarbons or hydrocarbons which cannot be vaporized in conventional commercial heaters without substantial decomposition. These heavy constituents form part of the residuum from ordinary distillations, hence the name residual. Thus the charge stock to the initial or absorption hydrodesulfurization stage may be, for example, crude oil or topped or reduced crude. The hydrosulfurized product from the first stage is topped to remove at least gasoline and lighter constituents and either the entire remainder, including the hydrodesulfurized residual constituents or a suitable cut therefrom, such as gas oil, may be charged-to the second or catalytic cracking stage.

The conditions-and catalysts employed in the cracking stage form no part of my invention. Consequently, it is considered unnecessary to list these in detail, particularly since these catalysts and conditions are well known to the art.

One principal advantage of my invention is in the reduced coke lay'dotvn on the cracking catalyst. This is accomplished while at the same time increasing the yield of cracked gasoline. The reduction of coke laydown is advantageous in view of the decreased regeneration requirements. Other advantages include production or the following: low sulfur cracked and straight run gasoline products without need of acid treating; high ASTM octane numbers, clear and leaded for straight run gasoline; reduced corrosion and maintenance costs on equipment; reduced sulfur content of cracked fuel gas and fuel oil low viscosity and low yield of tar; low sulfur content of the C4 fraction, and reduced regeneration requirements for both the cracking and hydrogenating catalysts.

What I'cl'aim is: V

l. A process for. treating a sulfur-containing hydrocarbon oil selected from the group consisting of crude oil, topped crude, and reduced crude, comprising contacting the oil'with a material selected from the group consisting of iron,- nickel, cobalt, their oxides and combinations thereof, in a first zone at elevated temperature and pressure and in the presence of added hydrogen, chemically absorbing sulfur in the oil on said material, terminating the contacting before substantial amounts of hydrogen sulfide occur in the product, topping the desulfurized product to remove at least gasoline and lighter constituents, catalytically cracking the remainder in a second zone isolated from the first zone, said second zone being maintained at cracking temperature and containing a cracking. catalyst incapable of chemically absorbing substantial amounts of sulfur,

said catalytic cracking being carried out to a depth of conversion of at least about 30 per cent.

2. A process for treating a sulfur-containing hydrocarbon oil selected from the group consisting of crude oil, topped crude and reduced crude, comprising contacting the oil-with a material selected from the group consisting of iron, nickel, cobalt, their oxides and combinations thereof at a temperature between about 750 and about 950 F. at a pressure between about and about 1500 p.- s. i. g. in the presence of added hydrogen, said contacting being carried cut at a space velocity between about 0.2' and about 6' liquid volumes of oil per volume ofcatalyst per liour'and at a hydrogen on ratio of between about 1600 andabout 20,000 cubic feet of hydrogen per barrel of oil, chemically absorbing sulfur in the oil on said material, terminating the contact n liefore substantial amounts of hydrogen sulfide occur in the product, topping the'desulfuriaed product to remove at least gasoline and lighter con= stituents, catalytically cracking the remainder in a second zone isolated from thefirst zone, said second zone being maintained at cracking temperature and containing a cracking catalyst ine capable of chemically absorbing substantial amounts of sulfur, said catalytic cracking being carried out to a depth of conversion of at least about 30 per cent.

3. A process for treating a sulfur-containing hydrocarbon oil selected from the group consist= ing of crude oil, toppedcrude and reduced crudecornprising contacting the oilwith a material selected from the group consisting of iron,- nickel; cobalt, their oxides and combinations theredf, in a first zone at a temperature between about 800 and about 870 F. and at a pressure between about 500 and about 1000 p. s. i. g. in the presence of added hydrogen, said contacting being carried out at a space velocity between about 0.5 and about 2.0 liquid volumes of oil per volume of catalyst per hour and at a hydrogen 3 oil ratio of between about 5000 and about 12,000 cubic feet of hydrogen per barrel of oil, chemically absorbing sulfur in the oil on said material, terminating the contacting before substantial amounts of hydrogen sulfide occur in the product. topping the desulfurized product to remove at least gasoline and lighter constituents, catalytic'ally cracking the remainder in a second zone isolatedfrofii the first zone, said second zonebeing maintained at cracking temperature and contairufig a cracking catalyst incapable of chemically-absorbing substantial amounts of sulfur,- said cataiyuc cracking being carried out to a depth'ofconversion of at least about 30 per cent.

WILLIAM- A. SLATER;

REFERENCES CITE'D The following references areof' record in the file of this patent:

UNITED STATES PATENTS Number Name Date 2.4.503724 Grote Oct; 5'; 1948 2,516,877 Home et al Aug. 1. 1950 

1. A PROCESS FOR TREATING A SULFUR-CONTAINING HYDROCARBON OIL SELECTED FROM THE GROUP CONSISTING OF CRUDE OIL, TOPPED CRUDE, AND REDUCED CRUDE, COMPRISING CONTACTING THE OIL WITH A MATERIAL SELECTED FROM THE GROUP CONSISTING OF IRON, NICKEL, COBALT, THEIR OXIDES AND A COMBINATION THEREOF, IN A FIRST ZONE AT ELEVATED TEMPERATURE AND PRESSURE AND IN THE PRESENCE OF ADDED HYDROGEN, CHEMICALLY ABSORBING SULFUR IN THE OIL ON SAID MATERIAL, TERMINATING THE CONTACTING BEFORE SUBSTANTIAL AMOUNTS OF HYDROGEN SULFIDE OCCUR I N THE PRODUCT, TOPPING THE DESULFURIZED PRODUCT TO REMOVE AT LEAST GASEOLINE AND LIGHTER CONSTITUENTS, CATALYTICALLY CRACKING THE REMAINDER IN A SECOND ZONE ISOLATED FROM THE FIRST ZONE, SAID SECOND ZONE BEING MAINTAINED AT CRACKING TEMPERATURE AND CONTAINED. 